Integrated process for the preparation of alkyl and alkenyl substituted aromatic compounds

ABSTRACT

Integrated process for the production of alkyl and alkenyl substituted aromatic compounds which comprises simultaneously dehydrogentaing in a reactor-regenerator system a mixture containing an alkyl and an aromatic alkyl hydrocarbon coming from an alkylation unit and recycling the dehydrogenated alkyl hydrocarbon thus produced, after separation, to the alkylation unit.

The present invention relates to an integrated process for thepreparation of alkyl and alkenyl substituted aromatic compounds.

More specifically, the present invention relates to an integratedprocess for the preparation of alkyl substituted aromatic compounds,such as ethylbenzene, and alkenyl substituted aromatic compounds, suchas styrene and α-methylstyrene (via cumene), from an aromaticderivative, such as benzene, and an alkane, such as ethane or propane.

Even more specifically, the present invention relates to an integratedprocess for the production of ethylbenzene and styrene with thecontemporaneous dehydrogenation of ethylbenzene, to give styrene, andethane, to give the ethylene necessary as reagent for the synthesis ofethylbenzene.

As is well known, styrene and α-methylstyrene are products which areused in the production of thermoplastic polymers, such as polystyrene,acrylonitrile-butadiene-styrene copolymers, styrene-acrylonitrileresins, styrene-butadiene elastomeric copolymers and in formulations forunsaturated polyester resins.

Styrene is generally prepared by the catalytic dehydrogenation ofethylbenzene by means of an adiabatic or isotherm system and in thepresence of catalysts selected from metallic oxides or their mixtures,whereas ethylbenzene is prepared by the alkylation of benzene, availableas a refinery product, with ethylene coming from cracking or from thedehydrogenation of ethane.

The alkylation reaction can be carried out in vapour phase, using ascatalysts, zeolites with high SiO₂/Al₂O₃ ratios, for example ZSM-5zeolites or Lewis acids, or in liquid phase. Details on the synthesis ofethylbenzene and on its dehydrogenation to produce styrene are providedin the Stanford Research Institute (SRI International) Reports.

The patent U.S. Pat. No. 6,031,143 describes an integrated process forthe production of ethylbenzene and styrene which comprises the followingoperating steps:

-   -   feeding a stream of benzene and a recycled stream containing        ethylene to an alkylation unit;    -   mixing the stream leaving the alkylation unit, containing        ethylbenzene, with a stream consisting of ethane;    -   feeding the mixture thus obtained to a dehydrogenation unit        containing a catalyst capable of simultaneously dehydrogenating        ethane and ethylbenzene to give ethylene and styrene        respectively;    -   feeding the product leaving the dehydrogenation unit to a        separation section to produce a stream essentially consisting of        styrene and a stream containing ethylene;    -   recycling the stream containing ethylene to the alkylation unit.

The dehydrogenation unit comprises a first fluid bed dehydrogenationreactor and a second regeneration reactor of the exhausted catalyst. Thelatter is continuously removed from the bottom of the first reactor andis fed to the head of the second reactor where it is kept under fluidconditions by pre-heated air which flows upward. In this way theexhausted solid slowly descends downwards in counter-current to the hotair which is rising and during this descent, it is regenerated, as thecarbonaceous residues which poison it are burnt. The passage of thecatalyst from one reactor to the other is guaranteed by a carrier gassuch as air or nitrogen.

The contemporaneous dehydrogenation of ethane and ethylbenzene, however,creates drawbacks as these two products have characteristics which makeit difficult to obtain acceptable conversions and selectivity toethylene and styrene, under the same operating conditions. For example,to obtain a 50% equilibrium conversion of ethylbenzene to styrene atatmospheric pressure, it is necessary to operate at a temperature ofabout 615° C. whereas under the same conditions, the equilibriumconversion of ethane to ethylene is only about 20%. To obtain a 50%equilibrium conversion of ethane to ethylene, it would be necessary tooperate at a minimum of 720° C., a temperature which would cause thethermal degradation of both the ethylbenzene and styrene. For moredetails, reference can be made to Paul H. Emmett“Catalysis-Hydrogenation and Dehydrogenation” vol. III, 453–471, 1995,Reinhold Publishing Corporation.

The operating conditions for embodying the process described in the U.S.patent cited above are therefore rather limited and require aparticularly controlled running of the dehydrogenation reactor.

The Applicant has now found an integrated process for the production ofalkyl substituted aromatic compounds, such as ethylbenzene, and alkenylsubstituted aromatic compounds, such as styrene, with a greateroperating flexibility and wider selection of catalyst which involves theuse of a fluid bed dehydrogenation reactor in which the feeding of thealkane (ethane) is at least partially differentiated with respect tothat of the ethylbenzene, as described below, exploiting the fact thatin a fluid bed reactor/regenerator system with the circulation of asolid there are different temperature zones. In fact, in the fluid bedunit of the reactor/regenerator system, the heat necessary fordehydrogenation is supplied by the hot regenerated catalyst which istransferred in continuous, by means of specific transport lines, fromthe regenerator, operating at a higher temperature, to thedehydrogenation reactor.

The object of the present invention therefore relates to an integratedprocess for the production of alkyl and alkenyl substituted aromaticcompounds, such as ethylbenzene and styrene, which comprises:

-   -   a) feeding to an alkylation unit, a stream consisting of a        C₆–C₁₂ aromatic hydrocarbon and a recycled stream containing a        C₂–C₅ alkenyl hydrocarbon;    -   b) optionally mixing the stream leaving the alkylation unit,        containing the alkylaromatic compound, with a stream consisting        of a C₂–C₅ alkyl hydrocarbon;    -   c) feeding the stream of step (b) to a fluid bed        dehydrogenation/regeneration unit containing a catalyst capable        of dehydrogenating, also simultaneously, the alkyl hydrocarbon,        optionally present, and the alkylaromatic compound;    -   d) continuously discharging the exhausted catalyst which        accumulates on the bottom of the dehydrogenation reactor and        feeding it to the head of the regeneration reactor;    -   e) continuously discharging the regenerated catalyst which        accumulates on the bottom of the regeneration reactor and        feeding it to the head of the dehydrogenation reactor;    -   f) feeding the hydrocarbon stream leaving the dehydrogenation        reactor to a separation section to produce a stream essentially        consisting of the alkenyl substituted aromatic compound and a        stream containing the alkenyl hydrocarbon;    -   g) recycling the stream containing the alkenyl hydrocarbon to        the alkylation unit;        said integrated process being characterized in that the fluid        for transporting the catalyst, which is deposited on the bottom        of the regenerator at a temperature of 650–800° C., to the        dehydrogenation reactor consists of a C₂–C₅ alkyl hydrocarbon.

According to the present invention, a first stream is fed to thealkylation unit, consisting of an aromatic hydrocarbon, for example astream of fresh refinery grade benzene charge, consequently having apurity higher than or equal to 95% by weight, and a second, recycledstream, essentially consisting of the alkenyl hydrocarbon, such asethylene, and non-converted alkyl hydrocarbon, such as ethane. Morespecifically, this second stream consists of 20–95% in moles, preferably40–85%, of ethane and 5–80% in moles, preferably 15–60% of ethylene,respectively.

In the recycled stream, 0.1–2% by weight (calculated with respect to thetotal ethylene+ethane weight) of other light products, for examplemethane and hydrogen, formed both in the alkylation phase anddehydrogenation phase, are also present.

The two streams are fed to the alkylation unit so as to havebenzene/ethylene molar ratios required by current technologies,typically between 1.8 and 50, preferably between 2 and 10. Thealkylation reaction is carried out with conventional systems, forexample according to the method described in European patent 432,814.

Any alkylation reactor can be used in the process, object of the presentinvention, such as fixed bed or fluid bed reactors, carrier bed reactorsand catalytic distillation reactors. For example, the catalyticdistillation reactor can be used, which operates in mixed gas-liquidphase, described in U.S. Pat. No. 5,476,978 and in publishedinternational patent application WO 98/09929. In a catalyticdistillation reactor, the reagents and catalytic reaction products, inthe present case the reagents and alkylation reaction products, aresimultaneously separated by distillation using the catalytic reactor asdistillation column.

The preferred alkylation catalysts comprise synthetic and natural porouscrystalline solids such as acid zeolites in which the atomic ratiosilicon/aluminum ranges from 5/1 to 200/1. In particular, Y, betazeolites, mordenite, omega, A, X and L zeolites or porous crystallinesolids MCM-22, MCM-36, MCM-49, MCM 56 and ERS-10, are preferred.

In an alternative embodiment of the present invention, the alkylationreaction can be carried out using a continuous fixed bed reactorfunctioning in gaseous phase described, for example, in U.S. Pat. Nos.4,409,412 and 5,517,184. In this case, the catalyst is selected fromzeolites of the ZSM group in which the atomic ratio silicon/aluminumranges from 20/1 to 200/1. Examples of ZSM-type zeolites are ZSM-5,ZSM-11, ZSM-12, ZSM-23, ZSM-35, ZSM-38 and ZSM-48 zeolites. ZSM-5 isparticularly preferred.

The alkylation reaction can be carried out under temperature andpressure conditions which depend, as is well known to experts in thefield, on the type of reactor and selection of reagents. In the case ofthe alkylation of benzene with ethylene, the reaction temperaturegenerally ranges from 50 to 450° C. More specifically, for processes ingas phase, the temperature preferably ranges from 325 to 450° C. whereasin the case of a catalytic distillation reactor operating in mixedgas-liquid phase, the reaction temperature, varying along the catalyticbed, ranges from 140 to 350° C., preferably from 200 to 300° C.

The pressure inside the alkylation reactor is kept at values rangingfrom 0.5 to 6 MPa, preferably from 2 to 4.5 MPa.

The aromatic stream leaving the alkylation reactor can be treated withthe conventional means to respectively obtain a substantially purestream of non-converted aromatic product, for example benzene, asubstantially pure stream of alkyl substituted aromatic compound, forexample ethylbenzene, and a stream of heavier products essentiallyconsisting of di- or polyalkyl substituted aromatic compounds, forexample di- or polyethylbenzenes.

The separation system preferably consists of a series of distillationcolumns, from the first of which non-reacted benzene is recovered andrecycled to the alkylation reactor or to a transalkylation unit asdescribed below. Ethylbenzene is recovered from the second distillationcolumn and fed to dehydrogenation, whereas polyethylbenzenes, such asdiethylbenzenes and triethylbenzenes are recovered from the thirdcolumn.

The polyalkyl substituted aromatic compounds, such as polyethylbenzenes,can be fed to a transalkylation reactor for transalkylation with C₆–C₁₂aromatic hydrocarbons, in the case in question with benzene, to producethe corresponding monoalkyl substituted aromatic compounds, such asethylbenzene, and increase the yield of the alkylation reaction.

The transalkylation reactor preferably consists of a fixed bed reactorfunctioning in liquid phase in which a conventional transalkylationcatalyst is present, such as Y zeolite, beta zeolite or mordenite,preferably Y zeolite or beta zeolite. The transalkylation reaction canbe carried out according to what is described in U.S. Pat. No.5,476,978.

In the case of the transalkylation of polyethylbenzenes with benzene,the benzene/ethylene molar ratio, calculated with respect to the totalmoles of benzene present as such and as polyethylbenzene and withrespect to the total moles of ethylene present as substituent in thepolyethylbenzenes, ranges from 1.8/1 to 17/1, preferably from 2.4/1 to10/1. The temperature in the transalkylation reactor is maintained at 50to 300° C., preferably from 120 to 250° C., whereas the pressure is keptat 0.02 to 5.5 MPa, preferably from 0.7 to 4.5 MPa.

The C₂–C₅ alkyl hydrocarbon or, in the preferred case, ethane which canbe optionally mixed with the alkylation product, is a stream of freshcharge deriving from refineries, and is therefore available, likebenzene, with a purity higher than or equal to 95% by weight. The ethanefed in this phase is generally equal to 0–70% by weight of the totalethane.

The stream containing the alkylation product, optionally mixed withethane, is fed in gas phase to the base of the dehydrogenation reactorwhich operates at a temperature ranging from 450 to 650° C. and at apressure ranging from 0.1 to 3 ata, preferably at atmospheric pressureor a slightly higher value, and with a flow-rate of the reagents,expressed as hourly volumetric flow-rate of the reagents per liter ofcatalyst (Gas Hourly Space Velocity or GHSV) ranging from 100 to 10,000h⁻¹, preferably from 100 to 1,000 h⁻¹, with a residence time of thecatalyst in the fluid bed zone ranging from 5 to 30 minutes, preferablyfrom 10 to 15 minutes.

To obtain an optimum dehydrogenation, the catalyst is charged into theupper part of the reactor and maintained in the fluid state by thehydrocarbon stream, fed to the base, so as to slowly descend towards thebottom in countercurrent to the gaseous phase which is rising. Duringthis descent, the catalyst is gradually deactivated and collects on thebottom substantially exhausted.

The exhausted catalyst is continuously removed from the bottom of thehydrogenation reactor and is fed, by means of a carrier fluid, such asair or nitrogen, to the regeneration reactor. The regeneration reactorsubstantially operates in the same way as the dehydrogenation reactor.The exhausted solid is charged into the upper part of the reactor and ismaintained in the fluid state by preheated air, optionally enriched withoxygen, so as to slowly descend towards the bottom in counter-currentwith the hot air which is rising. During this descent the carbonaceousresidues present on the catalyst are gradually burnt so that thesubstantially regenerated catalyst collects on the bottom of theregenerator. Owing to the high selectivity of the dehydrogenationreactions, it is also possible to feed fuel gas to the regenerator tosupply the necessary heat for completing the thermal balance of thesystem by its combustion.

In the regenerator, it is preferable to operate at atmospheric pressure,or slightly higher values, at a space velocity ranging from 100 to 1,000h⁻¹ and with residence times of the solid ranging from 5 to 60 minutes,preferably from 20 to 40 minutes. The temperature profile inside theregeneration reactor generally ranges from 600 to 800° C.

The regenerated catalyst, at a temperature of about 650–800° C., iscontinuously removed from the bottom of the regeneration reactor and isfed to the dehydrogenation reactor using the C₂–C₅ alkyl hydrocarbon orethane, as carrier fluid, in a quantity ranging from 30 to 100% byweight of the total used, preferably from 50 to 70%. During the transferfrom the regenerator to the dehydrogenation reactor, the ethane isconverted to ethylene, cooling the catalyst which is thus fed to thedehydrogenation reaction to create an optimum temperature profile in thereactor for the conversion of ethylbenzene to styrene.

Any catalyst capable of dehydrogenating, also simultaneously, a paraffinsuch as ethane and an alkylaromatic hydrocarbon such as ethylbenzene canbe used in the process, object of the present invention. For example, aparticularly suitable catalyst is that described in international patentapplication PCT/EP 00/9196 based on iron and one or more promoters,selected from alkaline or earth alkaline metals and lanthanides, onalumina in delta or theta phase or in a mixed delta+theta, theta+alphaor delta+theta+alpha phase, modified with silica, and having a surfacearea of preferably less than 150 m²/g, determined with the BET method.More specifically, it is a catalyst which comprises:

-   -   1–60% by weight, preferably 1–20%, of iron oxide;    -   0.1–20% by weight, preferably 0.5–10% of at least one alkaline        or earth alkaline metal oxide, for example potassium;    -   0–15% by weight, preferably 0.1–7% of a second promoter selected        from lanthanide oxides, for example cerium, lanthanum or        praseodymium;    -   the complement to 100 being alumina modified with 0.08–5% by        weight of silica.

Further examples of catalysts are those based on gallium and platinumdescribed in European patent 637,578 or based on chromium and tindescribed in European patent 894,781. Other dehydrogenation catalystsfor paraffins and/or alkylaromatic hydrocarbons are described inEuropean patents 400,448 and 335,130 and in international patentapplication WO 96/34843.

The catalyst based on gallium and platinum can be selected from thosecomprising:

-   -   0.1–34% by weight, preferably 0.2–3.8%, of Ga₂O₃;    -   1–99 ppm (by weight), preferably 3–80 ppm, of platinum;    -   0.05%–5% by weight, preferably 0.1–3%, of an alkaline and/or        earth alkaline oxide, for example potassium;    -   0.08–3% by weight of silica;        the complement to 100 being alumina in delta or theta phase or        in a mixture of delta+theta, theta+alpha or delta+theta+alpha        phases with a surface area of less than 150 m²/g, determined        with the BET method.

The catalyst based on chromium and tin can be selected from thosecomprising:

-   -   6–30% by weight, preferably 13–25%, of Cr₂O₃;    -   0.1–3.5% by weight, preferably 0.2–2.8%, of SnO;    -   0.4%–3% by weight, preferably 0.5–2.5%, of an alkaline oxide,        for example potassium;    -   0.08–3% by weight of silica;        the complement to 100 being alumina in delta or theta phase or        in a mixture of delta+theta, theta+alpha or delta+theta+alpha        phases with a surface area of less than 150 m²/g, determined        with the BET method.

At the end of the dehydrogenation, a dehydrogenated stream is recovered,essentially consisting of ethylene and styrene. More specifically, thestream comprises: 15–30% by weight of styrene; 7–15% by weight ofethylene; 10–50% by weight of non-reacted ethylbenzene and 15–55% byweight of non-reacted ethane, plus other products such as hydrogen,methane, toluene, benzene formed both during the alkylation phase andduring the dehydrogenation phase.

The dehydrogenated stream is cooled, filtered and sent to a distillationsection for the recovery of the styrene and non-reacted ethylbenzene,which is recycled to the dehydrogenation, and the recovery of the streamcontaining ethylene which is recycled, as feeding, to the alkylationunit.

If the dehydrogenation catalyst available is particularly active, i.e.low contact times of the reagent gas with the catalyst are necessary foreffecting the dehydrogenation reactions, the dehydrogenation reactor canbecome a reactor in equicurrent, in which the solid is completelycarried upwards pneumatically by the gas (riser-type reactor). In thiscase the superficial velocity of the gas must be higher than theterminal velocity of the largest particles present in the fluid bed. Thesuperficial velocity of the gas phase is therefore in the order of atleast a few m/s. The space velocity (GHSV) for this reactor is greaterthan 500 h⁻¹ and preferably greater than 1000 h⁻¹. In this case thealkyl hydrocarbon is fed to the bottom of the riser, entering intocontact with the catalyst at the maximum reaction temperature. Thestream containing the alkylaromatic compound is, on the other hand,injected at a suitable height along the riser when most of thedehydrogenation of the alkyl hydrocarbon has already taken place and thetemperature has dropped to levels compatible with the correctdehydrogenation reaction trend of the alkylaromatic compound.

The integrated process for the production of ethylbenzene and styrene,object of the present invention, can be better understood by referringto the block scheme of the enclosed figure which represents anillustrative but non-limiting embodiment.

With reference to the scheme, (A) represents the alkylation unit, (D)the dehydrogenation reactor, (R) the regeneration unit of the catalyst,(C) a water condenser, (S) a scrubber, (SP) a separation section bymeans of distillations in series, (F) a filtration unit, (G1) and (G2)represent two gas-gas heat exchangers, (K1) and (K2) are compressors,(V) a gas-liquid separator, (LT) a membrane separation unit, (T1) and(T2) are the pneumatic carrier lines of the catalyst between reactor andregenerator and (ST) the stack for discharging the fumes into theatmosphere.

The present invention is therefore clearly illustrated on the basis ofthe enclosed scheme and previous description. In fact, a stream (1)consisting of benzene and a recycled stream (14) essentially consistingof ethylene and ethane, together with traces of hydrogen and methane,are fed, as reagents, to the alkylation unit (A). The inert products(3), which would otherwise accumulate in the production cycle, areflushed from the alkylation unit.

The alkylated stream (4), essentially consisting of ethylbenzene andethane, is mixed to a second recycled stream (16), containingethylbenzene, coming from the distillation section (S). A part of theethane necessary for the integrated process, object of the presentinvention, can be mixed, by means of line (2), to the stream (4).

The mix (5) thus obtained, after preheating in (G1), is fed, by means ofline (7), to the dehydrogenation reactor (D). The reactor (D) operatestogether with the catalyst regeneration unit (R). In particular, theexhausted catalyst which accumulates on the bottom of (D) iscontinuously removed and pneumatically conveyed, through line (T1) andwith the introduction of carrier gas, for example, air or nitrogen, tothe upper part of the regenerator (R). The stream of air (21), takenfrom the atmosphere (19), compressed in (K2) to give stream (20) andwhich is preheated in (G2), is fed to the regenerator. The air (21), fedto the base by means of a suitable distributor, not illustrated in thefigure, burns the carbonaceous deposits (coke) deposited on the surfaceof the catalyst and, rising in countercurrent, keeps the solid influidized state. The effluent gases (22) from the regenerator are cooledin (G2), filtered in (F) and discharged from (ST).

Analogously, the regenerated catalyst, which accumulates on the bottomof (R), is continuously removed and pneumatically conveyed, through line(T2), using ethane (6) as carrier gas, to the upper part of thedehydrogenation reactor (D). During the transfer phase, the ethane isthoroughly mixed with the hot catalyst and is partially transformed toethylene, lowering the temperature of the catalyst to values compatiblewith the dehydrogenation of ethylbenzene.

The dehydrogenated product (8), which essentially consists of styrene,ethylene, non-converted ethylbenzene and ethane, methane, hydrogen andother products, such as toluene and benzene, is cooled in (G1), washedfrom the entrained powders in (S), further cooled in the condenser(C)and fed to the separator (V). A stream (12) of condensable products,essentially consisting of styrene, ethylbenzene and other by-products(benzene, toluene) is recovered from the bottom of (V) whereas a stream(11) of light products essentially consisting of ethylene, ethane,methane and hydrogen is recovered at the head.

The stream (12) goes to the distillation unit (S), for example a unitcomprising one or more distillation columns, from which high purity(>99.5%) styrene (18) is recovered together with ethylbenzene (16),recycled to the dehydrogenation, and by-products (17) which are sent forsubsequent treatment.

The stream (11) is brought to the operating pressure of the alkylationunit in (K1), separated from the hydrogen (15) in the membrane removalsystem (LT) and recycled to (A), as primary feed, by means of line (14).

An illustrative but non-limiting example is provided hereunder for abetter understanding of the present invention and for its embodiment.

EXAMPLE

An integrated plant is described, for the production of styrene, whichoperates for 8,400 hours/year with a normal yearly production of 3,500tons of styrene.

A contemporaneous dehydrogenation of ethane and ethylbenzene is effectedanalogously to the procedure described in U.S. Pat. No. 6,031,143. Theethylbenzene necessary for the production of styrene is premixed withethane so that the feeding to the reactor consists of 30% molar ofethylbenzene and 70% molar of ethane. The reaction is carried out at anaverage pressure in the fluid bed of 1.5 atmospheres and at atemperature ranging from 550° C. at the bottom of the reactor to 600° C.at the upper end of the catalytic bed, where the hot regeneratedcatalyst coming from the reactor, is fed. The space velocity (GHSV) is300 Nl/h of gas per liter of catalyst. The dehydrogenation catalystcomprises gallium oxide (2.33% by weight), potassium oxide (0.6% byweight), platinum (75 ppm), silica (1.56% by weight), the complement to100 being alumina, and the residence time of the solid in the reactor isequal to 12 minutes. The ethylbenzene conversion is 52% by weight andthe selectivity to styrene 92% by weight. The ethane conversion is 10%by weight and the selectivity to ethylene 90% by weight. In this way,the molar ratio between reacted ethylbenzene and ethylene produced isequal to 2.5.

Another amount of ethane, equal to 60% of the quantity premixed withethylbenzene, is fed to the base of the carrier line which brings theregenerated catalyst to an average temperature of 650° C. and to anaverage pressure of 2 bars from the regenerator to the reactor.

The ethane acts as carrier gas but also partly reacts to form ethylene.The yield to ethylene is 20% by weight, and consequently, after theeffluent gas from the fluid bed of the reactor is mixed with the carriergas from the regenerator to the reactor, the molar ratio between reactedethylbenzene and ethylene formed is equal to 0.99. A quantity ofethylene was therefore produced, by the dehydrogenation of ethane, whichwas sufficient to be used as reagent in the alkylation section andproduce all the ethylbenzene which reacts in the dehydrogenationreactor.

1. An integrated process for the production of alkyl and alkenylsubstituted aromatic compounds which comprises: a) feeding to analkylation unit, a stream consisting of a C₆–C₁₂ aromatic hydrocarbonand a recycled stream containing a C₂–C₅ alkenyl hydrocarbon; b)optionally mixing the stream leaving the alkylation unit, containing thealkylaromatic compound, with a stream consisting of a C₂–C₅ alkylhydrocarbon; c) feeding the stream of step (b) to a fluid beddehydrogenation/regeneration unit containing a catalyst capable ofdehydrogenating, also simultaneously, the alkyl hydrocarbon, optionallypresent, and the alkylaromatic compound; d) continuously discharging theexhausted catalyst which accumulates on the bottom of thedehydrogenation reactor and feeding it to the head of the regenerationreactor; e) continuously discharging the regenerated catalyst whichaccumulates on the bottom of the regeneration reactor and feeding it tothe head of the dehydrogenation reactor; f) feeding the hydrocarbonstream leaving the dehydrogenation reactor to a separation section toproduce a stream essentially consisting of the alkenyl substitutedaromatic compound and a stream containing the alkenyl hydrocarbon; g)recycling the stream containing the alkenyl hydrocarbon to thealkylation unit; said integrated process being characterized in that thefluid for transporting the catalyst, which is deposited on the bottom ofthe regenerator at a temperature of 650–800° C., to the dehydrogenationreactor consists of a C₂–C₅ alkylhydrocarbon.
 2. The integrated processaccording to claim 1, wherein the C₆–C₁₂ aromatic hydrocarbon isbenzene.
 3. The integrated process according to claim 1, wherein theC₂–C₅ alkyl/alkenyl hydrocarbon is ethane/ethylene.
 4. The integratedprocess according to claim 1, wherein the recycled stream consists of20–95% in moles of ethane and 5–80% in moles of ethylene, respectively.5. The integrated process according to claim 1, wherein the streams arefed to the alkylation unit so as to have molar ratios benzene/ethyleneranging from 1.8 to
 50. 6. The integrated process according to claim 1,wherein the alkylation unit comprises a catalytic distillation reactorwhich operates in mixed gas-liquid phase.
 7. The integrated processaccording to claim 1, wherein the alkylation unit comprises a continuousfixed bed reactor functioning in gas phase.
 8. The integrated processaccording to claim 6, wherein the alkylation catalyst is selected fromsynthetic and natural porous crystalline solids wherein the atomic ratioof silicon/aluminum ranges from 5/1 to 200/1.
 9. The integrated processaccording to claim 8, wherein the alkylation catalyst is selected fromthe group consisting of Y zeolites, beta zeolites, mordenite zeolites,omega zeolites, A zeolites, X zeolites, L zeolites, porous crystallinesolid MCM-22, porous crystalline solid MCM-36, porous crystalline solidMCM-49, porous crystalline solid MCM 56 and porous crystalline solidERS-10.
 10. The integrated process according to claim 7, wherein thealkylation catalyst is selected from zeolites of the ZSM group in whichthe atomic ratio silicon/aluminum ranges from 20/1 to 200/1.
 11. Theintegrated process according to claim 10, wherein the alkylationcatalyst is selected from ZSM-5, ZSM-11, ZSM-12, ZSM-23, ZSM-35, ZSM-38and ZSM-48 zeolites.
 12. The integrated process according to claim 11,wherein the alkylation catalyst is ZSM-5.
 13. The integrated processaccording to claim 1, wherein the alkylation reaction is carried out ata temperature ranging from 50 to 450° C.
 14. The integrated processaccording to claim 6, wherein the alkylation temperature, varying alongthe catalytic bed, ranges from 140 to 350° C.
 15. The integrated processaccording to claim 7, wherein the alkylation temperature ranges from 325to 450° C.
 16. The integrated process according to claim 1, wherein thepressure inside the alkylation reactor is maintained at values rangingfrom 0.5 to 6 MPa.
 17. The integrated process according to claim 1,wherein the alkylation unit comprises a separation system torespectively obtain a substantially pure stream of non-convertedaromatic hydrocarbon, a substantially pure stream of alkyl substitutedaromatic compound, and a stream of heavier products essentiallyconsisting of di- or polyalkyl substituted aromatic compounds.
 18. Theintegrated process according to claim 17, wherein the separation systemconsists of a series of distillation columns, from the first of whichnon-converted aromatic hydrocarbon is recovered and recycled to thealkylation reactor or to a transalkylation unit, alkyl substitutedaromatic compound is recovered from the second distillation column andfed to the dehydrogenation unit, and polyalkyl substituted aromaticcompounds are recovered from the third column.
 19. The integratedprocess according to claim 18, wherein the polyalkyl substitutedaromatic compounds are fed to the transalkylation unit fortransalkylation with C₆–C₁₂ aromatic hydrocarbons to producecorresponding alkyl substituted aromatic compounds.
 20. The integratedprocess according to claim 19, wherein the transalkylation is carriedout in a fixed bed reactor functioning in liquid phase in which atransalkylation catalyst selected from the group consisting of Yzeolite, beta zeolite and mordenite, is present.
 21. The integratedprocess according to claim 19, wherein, in the case of thetransalkylation of polyethylbenzenes with benzene, the molar ratiobenzene/ethylene calculated with respect to the total moles of benzenepresent as such and as polyethylene and on the total moles of ethylenepresent as substituent in the polyethylbenzenes, ranges from 1.8/1 to17/1.
 22. The integrated process according to claim 21, wherein thetemperature in the transalkylation reactor is maintained at 50 to 300°C. whereas the pressure is maintained at 0.02 to 5.5 MPa.
 23. Theintegrated process according to claim 1, wherein the C₂–C₅ alkylhydrocarbon fed to the stream leaving the alkylation unit is equal to0–70% by weight of the total.
 24. The integrated process according toclaim 1, wherein the stream containing the alkylation product is fed ingas phase to the base of the dehydrogenation reactor operating at atemperature ranging from 450 to 650° C. and at a pressure ranging from0.1 to 3 ata.
 25. The integrated process according to claim 1, whereinthe regeneration reactor is fed with preheated air and optionally withfuel gas to supply heat by its combustion.
 26. The integrated processaccording to claim 1, wherein the temperature profile inside theregeneration reactor generally ranges from 600 to 800° C.
 27. Theintegrated process according to claim 1, wherein the regeneratedcatalyst is continuously removed from the bottom of the regenerationreactor and is fed to the dehydrogenation reactor using the C₂–C₅ alkylhydrocarbon as carrier fluid in a quantity ranging from 30 to 100% byweight with respect to the total weight used.
 28. The integrated processaccording to claim 1, wherein the dehydrogenation catalyst is based oniron and one or more promoters, selected from alkaline or earth-alkalinemetals and lanthanides, on alumina in delta or theta phase or in amixture of delta+theta, theta+alpha or delta+theta+alpha phases,modified with silica, and having a surface area of preferably less than150 m²/g, determined with the BET method.
 29. The integrated processaccording to claim 28, wherein the dehydrogenation catalyst comprises:1–60% by weight, of iron oxide; 0.1–20% by weight of at least onealkaline or earth alkaline metal oxide; 0–15% by weight of a secondpromoter selected from lanthanide oxides; the complement to 100 beingalumina modified with 0.08–5% by weight of silica.
 30. The integratedprocess according to claim 1, wherein the dehydrogenation catalyst isselected from those based on gallium and platinum or based on chromiumand tin.
 31. The integrated process according to claim 30, wherein thecatalyst based on gallium and platinum comprises: 0.1–34% by weight ofGa₂O₃; 1–99 ppm (by weight) of platinum; 0.05%–5% by weight of analkaline and/or earth alkaline oxide; 0.08–3% by weight of silica; thecomplement to 100 being alumina in delta or theta phase or in a mixtureof delta+theta, theta+alpha or delta+theta+alpha phases with a surfacearea of less than 150 m²/g, determined with the BET method.
 32. Theintegrated process according to claim 30, wherein the catalyst based onchromium and tin comprises: 6–30% by weight of Cr₂O₃; 0.1–3.5% by weightof SnO; 0.4%–3% by weight of an alkaline oxide; 0.08–3% by weight ofsilica; the complement to 100 being alumina in delta or theta phase orin a mixture of delta+theta, theta+alpha or delta+theta+alpha phaseswith a surface area of less than 150 m²/g, determined with the BETmethod.
 33. The integrated process according to claim 1, wherein, at theend of the dehydrogenation, a dehydrogenated stream is recovered, whichcomprises 15–30% by weight of styrene; 7–15% by weight of ethylene;10–50% by weight of non-reacted ethylbenzene and 15–55% by weight ofnon-reacted ethane.
 34. An integrated process for the production ofalkyl and alkenyl substituted aromatic compounds which comprises: a)feeding to an alkylation unit, a stream consisting of a C₆–C₁₂ aromatichydrocarbon and a recycled stream containing a C₂–C₅ alkenylhydrocarbon; b) optionally mixing the stream leaving the alkylationunit, containing the alkylaromatic compound, with a stream consisting ofa C₂–C₅ alkyl hydrocarbon; c) feeding the stream of step (b) to a fluidbed dehydrogenation/regeneration unit containing a catalyst capable ofdehydrogenating, also simultaneously, the alkyl hydrocarbon, optionallypresent, and the alkylaromatic compound; d) continuously discharging theexhausted catalyst which accumulates on the bottom of thedehydrogenation reactor and feeding it to the head of the regenerationreactor; e) continuously discharging the regenerated catalyst whichaccumulates on the bottom of the regeneration reactor and feeding it tothe head of the dehydrogenation reactor; f) feeding the hydrocarbonstream leaving the dehydrogenation reactor to a separation section toproduce a stream essentially consisting of the alkenyl substitutedaromatic compound and a stream containing the alkenyl hydrocarbon; g)recycling the stream containing the alkenyl hydrocarbon to thealkylation unit; said integrated process being characterized in that thefluid for transporting the catalyst, which is deposited on the bottom ofthe regenerator at a temperature of 650–800° C., to the dehydrogenationreactor, consists of a C₂–C₅ alkyl hydrocarbon and that thedehydrogenation reactor is a riser reactor in equicurrent, in which thesolid is completely carried upwards pneumatically by a gas.
 35. Theintegrated process according to claim 34, wherein the space velocity ofthe gas (GHSV) is greater than 500 h⁻¹.
 36. The integrated processaccording to claim 34, wherein the alkyl hydrocarbon is fed to thebottom of the riser, entering into contact with the catalyst at themaximum reaction temperature, whereas the stream containing thealkylaromatic compound is fed at an intermediate height of the riserwhere the temperature has dropped to compatible levels for thecorresponding dehydrogenation.